EA035883B1 - Production of tert-butyl esters of ethylenically unsaturated carboxylic acids - Google Patents

Abstract

A process for continuously producing the tert-butyl ester of an ethylenically unsaturated carboxylic acid comprising the following steps: a) an ethylenically unsaturated carboxylic acid is reacted with isobutene in the presence of an acidic catalyst to give an esterification mixture; b) the acidic catalyst is removed; c) low-boiling-point constituents are removed; and d) a liquid containing tert-butyl ester is introduced into a distillation device and the product is purified by distillation in the distillation device, where d1) the liquid containing tert-butyl ester is separated in the distillation device to give a gaseous overhead product containing tert-butyl ester and a liquid bottom product containing carboxylic acid; d2) the gaseous overhead product containing tert-butyl ester is at least to some extent condensed and some of the condensate is returned to the distillation device; d3) the liquid bottom product containing carboxylic acid is at least to some extent returned to step a); d4) liquid bottom product containing carboxylic acid is drawn off and conducted to a heater, a superheated liquid return stream is taken from the heater and the superheated return stream is depressurized into the distillation device; and d5) at least in the overhead region of the distillation device those walls of the distillation device that are in contact with the vapours are thermally insulated and/or heated at least in some regions. The separation of the tert-butyl ester from unreacted carboxylic acid in the process proceeds with a particularly small extent of polymerization both of the tert-butyl ester and of the carboxylic acid.

Description

The present invention relates to a process for the continuous production of ethylenically unsaturated carboxylic acid tert-butyl ester by reacting a carboxylic acid with isobutene.

The tert-butyl esters of ethylenically unsaturated carboxylic acids can be used in various ways. Tert-butyl (meth) acrylates, for example, are important starting materials for the production of polymers, which are used, inter alia, as a component of paints, adhesives or paint resins. Tert-butyl esters of this type are usually obtained by acid-catalyzed addition of carboxylic acid to isobutene (Houben-Weyl, Methoden der Organischen Chemie [Methods of Organic Chemistry], vol. 8, 1952, p. 534; US 3031495 and US 3082246). The catalysts used are acids that are soluble in the reaction mixture, for example mineral acids or alkyl or aryl sulfonic acids (DE-A-1249857, US 3087962, US 3088969), or insoluble catalysts such as acidic ion exchange resins (US 3037052, US 3031495, DE-A-3105399, EP-A-268999).

The document WO 02/10109 A1 discloses a method for continuously producing tert-butyl ester of aliphatic C1-C ₄ carboxylic acid by reacting the carboxylic acid with isobutene in the liquid phase in the presence of an acid catalyst in a reactor, wherein the selection is performed ester by distillation from the reaction mixture obtained after removing unreacted isobutene and components with a low boiling point.

WO 02/10110 A2 describes a process for the preparation of tert-alkyl (meth) acrylate by reacting (meth) acrylic acid with an olefin in a homogeneous phase in the presence of an acid catalyst, wherein the tert-alkyl (meth) acrylate is isolated by distillation from the reaction mixture obtained after removing unreacted isobutene and components with a low boiling point.

Ethylenically unsaturated carboxylic acids and their esters can show a strong tendency to polymerize, in particular at elevated temperatures. These compounds are often exposed to high temperatures, in particular during the distillation process, which can lead to unwanted polymerization. This polymerization results in fouling of the device, clogging of pipelines and pumps, and accumulation of deposits on the column plates and the surfaces of the heat exchanger. Cleaning of plants is a labor-intensive and costly process that is harmful to the environment, in addition, due to the need for cleaning, the capacity utilization of the plant decreases. In addition, uncontrolled free radical polymerization is a safety risk.

Thus, it is an object of the present invention to provide a process for the continuous production of ethylenically unsaturated carboxylic acid tert-butyl ester, which separates the tert-butyl ester from unreacted carboxylic acid without polymerizing the tert-butyl ester and the carboxylic acid if possible. acid.

Document WO 2011/110257 A2 describes a method for producing a compound ready for polymerization, such as (meth) acrylic acid, from a liquid composition by distillation in a distillation unit, whereby a recirculation stream is withdrawn in the lower part of the distillation unit, which is in the form of a superheated liquid recirculation stream served in a distiller. In this method, the specified compound, ready for polymerization, is separated from solvents, for example.

WO 2011/110257 A2 does not describe the separation of ethylenically unsaturated ester from unreacted ethylenically unsaturated carboxylic acid. This distillation separation of the two ready-to-polymerize compounds is more complex, since it is necessary to ensure sufficient suppression of polymerization of the compositions obtained during the separation throughout the distillation unit. In addition, compared to lower esters, tert-butyl esters have a relatively high boiling point, which is a factor that increases the heat load during distillation.

DE 102008002923 A1 describes a method for producing tertiary alkyl esters of (meth) acrylic acid from (meth) acrylic acid and olefin, in which the (distillative) separation of the ester from (meth) acrylic acid is carried out with stabilization using a Noxyl compound.

DE 19539295 A1 describes a method for continuous distillation separation of liquid mixtures, the main component of which is (meth) acrylic acid, in a still. The substream of liquid supplied to the distiller is removed, and such a substream is fed to the distiller in a superheated form. The purified (meth) acrylic acid is removed through the top of the distiller.

Uses for tert-butyl (meth) acrylate copolymers include crosslinked paint formulations that are used, for example, as automotive paints or auto refinish paints. For this application, it is desirable that the acid content is low.

The object of the present invention is achieved by a method for the continuous production of ethylenically unsaturated carboxylic acid tert-1 035883 butyl ester, comprising the following steps:

a) the reaction of ethylenically unsaturated carboxylic acid with isobutene in the presence of an acid catalyst to obtain an esterification mixture;

b) removing the acid catalyst;

c) removal of low boiling point components; and

d) feeding the liquid containing tert-butyl ether to a distiller, where it is purified by distillation, while:

d1) in the still, the liquid containing tert-butyl ether is separated into a gaseous overhead product containing tert-butyl ether and a liquid underflow containing carboxylic acid;

d2) at least partial condensation of the gaseous overhead product containing tert-butyl ether occurs and at least partial recirculation of the condensate to the distiller as a reflux stream;

d3) at least partial recirculation of the liquid underflow containing carboxylic acid to step a);

d4) the carboxylic acid-containing liquid underflow is withdrawn and fed to a heating device; the superheated liquid recirculation stream is withdrawn from the heating device, which is released into the distiller; and d5) at least in the upper part of the distiller, the walls in contact with the steam are heated at least at some points, and / or such walls are thermally insulated at least at some points.

By the term underflow is meant a liquid product that is withdrawn at the bottom of the distiller, for example a liquid phase that is withdrawn through the bottommost part of the distiller or which is withdrawn from one to three trays located directly above this point. The term overhead means a gaseous or liquid product that is withdrawn at the top of the still, for example a liquid phase that is withdrawn through the topmost part of the still, or which is withdrawn from one to three trays immediately below the top.

During the heating and recirculation of the removed liquid underflow in accordance with the invention, the formation of a headspace is prevented in the entire line from the heater to the heater. Accordingly, the formation of gas phases which can condense in the line and form unstable liquid phases prone to polymerization are largely prevented.

In accordance with the method according to the invention, at least in the upper part of the distiller, the walls in contact with the steam are thermally insulated at least at some points and / or at least at some points are heated. As a result, the wall temperature in this region is preferably kept above the dew point of the tert-butyl ether. This prevents vapor condensation on the walls, which can lead to the formation of unstable liquid phases prone to polymerization.

The term top of the still refers to the region free of internals that is located above the topmost tray or above the topmost packing layer. It is usually formed by a domed body (a dome, for example, convex or toroidal spherical), which forms an element that closes the distiller from above.

The term "dew point of a compound" refers to the temperature at which, at a specified pressure, the compound condenses, i.e. the compound changes from gaseous to liquid state.

During esterification, an ethylenically unsaturated carboxylic acid reacts with isobutene in the presence of an acid catalyst to form an esterification mixture. The ethylenically unsaturated carboxylic acid is preferably selected from acrylic, methacrylic, dimethacrylic, ethacrynic, allylucetic, vinylacetic and vinylpropionic acid. In a preferred embodiment, the carboxylic acid is acrylic or methacrylic acid, with methacrylic acid being particularly preferred.

In a preferred embodiment, the acetic acid content of the ethylenically unsaturated carboxylic acid is less than 300 ppm, more preferably less than 100 ppm, particularly preferably less than 20 ppm, and the propionic acid content is less than 300 ppm, more preferably less than 200 ppm, particularly preferably less than 130 ppm. Since the removal of acetic and propionic acids from tert-butyl ester at further stages of the process will require high costs and efforts, if the goal is to obtain ethylenically unsaturated tert-butyl ester with a certain content of acetic and / or propionic acids, it is preferable to use ethylenically unsaturated carboxylic acid with controlled content of acetic and / or propionic acids. The content of acetic acid and propionic acid is usually determined by gas chromatography.

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Esterification is usually carried out in the liquid phase in the absence of a solvent. Therefore, catalysts are used which are at least partially soluble in the reaction mixture. Suitable catalysts are strong inorganic or organic acids. Strong inorganic acids are, for example, mineral acids such as sulfuric acid, phosphoric acid and polyphosphoric acid, preferably sulfuric acid. Strong organic acids are, for example, sulfonic acids such as p-toluene, benzene, dodecylbenzene and methanesulfonic acid, preferably p-toluenesulfonic and methanesulfonic acid. Inorganic catalysts, in particular, are only partially soluble in the reaction mixture at the beginning of the reaction. In the course of the reaction, the solubility of the catalyst increases (mainly due to the formation of a partial ester of the catalyst, for example, sulfuric acid monoester). Therefore, at least in the last section, the catalyst is typically present as a solution in the reaction mixture.

The concentration of catalyst in the esterification mixture is generally 0.1-10% by weight, preferably 0.5-5% by weight, based on the total amount of the esterification mixture.

The reaction of ethylenically unsaturated carboxylic acid with isobutene in the presence of an acid catalyst is preferably carried out in conventional reaction vessels or in columns (DE-A1128428). A suitable reactor is described, for example, in WO 02/10109 A1.

Preferably, the reaction is carried out in a reactor, which is in particular a cylindrical reactor. The reactor is divided into several separation sections, preferably 3, 4 or 5 separation sections. The sections are separated from each other by partitions that run at right angles to the longitudinal axis of the reactor. Each of these baffles has at least one opening so that the reaction mixture can pass from one section of the reactor to the next. The number of holes in the baffle is determined by the size of the reactor. Preferably, the baffles have only one opening, which, in particular, can be located in the middle of the baffle. The total area of the openings in the baffle is approximately 1 / 2000-1 / 500 of the cross-sectional area of the reactor.

The reactor sections can have the same or different volume. Preferably, the volume of the first section of the reactor is greater than the volume of the remaining sections. If the reactor has four sections, the following ratios of the individual sections in the total reactor volume have been found to be preferred:

Reactor section 1: 25% - 50%

Reactor section 2: 10% -25%

Reactor section 3: 10% -25%

Reactor section 4: 25% - 50%

Preferably, the reactor sections can be equipped with internals to improve mixing of the reaction volume. Suitable such internals are static mixing members and similar internals such as grids, distribution or sieve trays. It is especially preferred if internals of this type are located in the first section of the reactor and then in the upper half of the reactor.

The carboxylic acid is fed to the first section of the reactor, in particular to the zone at the base of the reactor, in liquid form. The acid supply can be carried out directly, for example by means of an immersion tube, but it is preferable if it is carried out by means of a device that ensures uniform distribution and mixing of the raw material. Devices of this type, such as distribution trays, perforated plates and pipes, nozzles, etc., are known to those skilled in the art. The carboxylic acid is preferably fed through a nozzle, thereby mixing the gas and liquid and mixing the contents of the reactor. Preferably, the nozzle is located at the base of the reactor. Suitable nozzle types (jet nozzle, mixing nozzle, two-phase nozzle, etc.) are known to those skilled in the art, which are also described, for example, in Ullmann's Encyclopedia of Industrial Chemistry, vol. B4, 5th ed., 1992, p. 280. In particular, when this type of nozzle is used, the flow in the first two reactor sections is turbulent and in the downstream sections it is laminar. Thus, in successive reaction sections, different flow characteristics are provided, for example, the flow can be turbulent, with a high level of backmixing, as in a mixing tank, or laminar with a low level of backmixing, as in a tubular reactor, thereby allowing a particularly preferred configuration the corresponding reaction section.

The catalyst is fed in the form of a mixture with a carboxylic acid, whereupon fresh catalyst or regenerated catalyst or a mixture thereof can be used.

It has been found that it is advantageous if at least a portion of the high boiling point liquid phase obtained in the catalyst removal step described below and / or at least a portion of the bottoms product obtained from the distillation purification is fed to the reactor. In this way, most of the acid catalyst and unreacted carboxylic acid are recycled.

Isobutene can be supplied in liquid and / or gaseous forms. The isobutene is preferably fed through a circular tube which has a plurality of outlets.

A portion of the reaction mixture can be withdrawn from the first and / or second sections of the reactor and recycled back to the corresponding section. This provides an increased mixing efficiency of the reaction mixture. Accordingly, the substream is recirculated through the aforementioned mixing nozzle into the first section of the reactor and / or through another nozzle in the region of the opening in the partition into the second section of the reactor. Such a further nozzle can be any of the types mentioned above for the mixing nozzle. A conical nozzle is preferably used. Such a nozzle is preferably positioned such that its outlet is approximately at the level of the partition separating the first section from the second. To control the temperature (in a closed loop), a certain substream that is taken out of the section can be passed through a heat exchanger.

The resulting esterification mixture is withdrawn through the top of the reactor and fed for further processing. Unreacted isobutene gas is collected at the top of the reactor. Preferably, condensable organic compounds, such as unreacted carboxylic acid, are condensed and separated from the isobutene-containing gas stream which is withdrawn through the top of the reactor, thereby removing gases inert to the esterification, such as air and butane. There is a partial dissolution of unreacted isobutene in the components that were separated by condensation. Then, the condensed organic compounds in liquid form are fed into the first section of the reactor, for example, through a mixing nozzle.

In general, the esterification temperature is in the range of about 10-40 ° C. Preferably, the esterification temperature is controlled so that the highest temperature is in the first section of the reactor. Preferably, the temperature in the first section of the reactor is in the range of about 30-40 ° C. The temperature in the second section is lower, preferably about 5-15 ° C. The temperature in the further sections after the second section may be the same or may be lower. As a rule, the temperature in them does not exceed the temperature in the second section, preferably in them the temperature is lower, in particular, it can be approximately 3-10 ° C lower. The temperature in the fourth section is generally the same as the temperature in the third section, or about 1-5 ° C lower. The temperature in the last section of the reactor is preferably in the range of 10-25 ° C.

The temperature distribution in the 4-section reactor is preferably as follows:

1st section: 33 - 38 ° С

2nd section: 23 - 28 ° С

3rd section: 15-22 ° С

4th section: 15 - 22 ° С

The temperature in the 3rd and 4th sections can be the same or different.

Since the addition of carboxylic acids to isobutene occurs with the release of heat, it is necessary to control the reaction temperature by removing the heat of reaction, in particular, in the first two sections of the reactor. In particular, this can be done by means of heat exchange devices, which can be configured as an external or internal heat exchanger. Cooling of the reactor walls can also be carried out. It has been found that it is preferable to carry out temperature control in the first two sections of the reactor using external heat exchangers, through which the feed of the reaction mixture is supplied and recirculated from a certain section of the reactor.

The esterification can be carried out under reduced pressure, under normal pressure, or under slightly elevated pressure (100-300 mbar abs.), Or preferably under elevated pressure (eg 0.5-3 bar).

The reaction mixture leaving the reactor contains a large proportion of the target ether. In addition, it contains unreacted reactants, catalyst, stabilizer, catalyst acid esters and other by-products in minor amounts. The reaction mixture contains only very small amounts of isobutene oligomerization product, typically <2 wt% based on the total amount of the reaction mixture.

First of all, the acid catalyst is removed from the esterification mixture. To remove the acid catalyst, it is preferable to partially evaporate the esterification mixture to obtain a high boiling liquid phase containing the acid catalyst and a vapor containing tert-butyl ether and isobutene. The high boiling liquid phase is generally at least partially recycled to the reactor.

Partial evaporation can be carried out in any way, but it is preferred if it is carried out in two stages. The evaporation is usually carried out at elevated temperature and under reduced pressure. Evaporation conditions are determined based on which product is the target.

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They are usually selected so that the temperature is in the range of approximately 50-150 ° C. The pressure is adjusted so that evaporation occurs quickly and gradually. For example, the pressure may be in the range of 10-200 mbar abs, more preferably in the range 30-90 mbar abs, most preferably in the range 50-70 mbar abs.

Any vacuum pumps can be used to create reduced pressure. It has been found that it is advisable to use pumps that do not use lubricating oil in order to avoid contamination. Especially preferred are oil-free twin-rotor vacuum pumps, which are referred to as screw vacuum pumps capable of operating without lubrication. Alternatively, liquid ring pumps can be used, in which, for example, the target ether acts as a barrier liquid.

The two-stage evaporation is preferably carried out in such a way that in the first stage 40-95% by weight, preferably 60-90% by weight, of the target ester is evaporated. In addition to tert-butyl ester and carboxylic acid, the vapor contains low boiling point components such as tert-butanol, tert-butyl acetate and diisobutene. The lower components obtained during the first distillation contain, as the first phase with a high boiling point, predominantly residual tert-butyl ether, carboxylic acid, acid catalyst and components with a high boiling point, for example polymeric (meth) acrylic compounds in the case of using (meth) acrylic acid ... 10-100 wt.% Of the first phase with a high boiling point is fed to the second stage of evaporation. If only a portion of the first high boiling phase is fed into the second evaporation stage, the remainder is recycled to the reactor. In the second stage of evaporation, the residual quantities of the target ester and most of the carboxylic acid (up to about 90 wt%) are evaporated.

The bottom stream obtained in the second stage of evaporation in the form of the second phase with a high boiling point contains predominantly acid catalyst, residual amounts of carboxylic acid and components with a high boiling point. At least a partial discharge, preferably a complete discharge of the second high boiling phase is carried out. However, it can also be recycled back to the reactor. The vapors obtained in both stages are combined and condensed. Typically, the distillate contains <20 ppm, in particular <10 ppm, catalyst.

Both evaporation steps can be carried out using conventional devices. However, it is preferable to use devices that provide fast distillation, such as film evaporators. Suitable film evaporators are known to those skilled in the art, see, for example, Ullmann's Encyclopedia of Industrial Chemistry, 5th ed., Vol. B3, 2-21-2-24 and 3-1-3-25, 1988 ...

Preferably, falling film or downdraft evaporators are used in the first step and thin film scraper or coiled tube evaporators are used in the second step.

It has been found that it is particularly advantageous in the first step to use a falling film evaporator as described, for example, in WO 02/10110.

The vapor, which predominantly contains the target ester, carboxylic acid and low boiling point components such as tert-butanol, tert-butyl acetate and diisobutene, is condensed by conventional means. Condensation is preferably carried out in one or more plate condensers or shell and tube condensers arranged in series. Preferably, shell and tube condensers are used with vertical tubes through which the steam flows from top to bottom. The non-condensed steam is preferably recycled to the reactor.

In accordance with one embodiment of the invention, vapor condensation occurs in two condensation stages, during which most of the ether is condensed at a first temperature and the remainder of the ether is condensed at a second temperature that is lower than the first temperature. As a result, there is an almost complete removal of tert-butyl ether, while it is accompanied by a minimum co-condensation of unreacted isobutene. Fractional condensation preferably takes place in two series-connected capacitors.

In order to avoid the deposition of solids on the contact surfaces of the condenser, an appropriate circulation of the overflow of the collected condensate is carried out to continuously flush the contact surfaces. The condensate that is pumped into the circulation stream can be distributed evenly, for example by means of a distributor, between the tubes of the shell-and-tube condenser while flowing down the inner walls of the condenser tubes. In case polymerizable carboxylic acids are used, the distribution of the stabilizer, described in more detail below, is achieved in the same way.

The condensate (combined) obtained from the steam condensation step is fed to a low boiling point removal step, which is preferably carried out by distillation. In the step of distillative removal of the low boiling point components, the low boiling point components, i. E. components other than isobutene, which have a lower boiling point than the target ester. Removal of the low boiling point components also provides additional amounts of unreacted isobutene, which are preferably separated from the low boiling point components and recycled to an appropriate stage. The low boiling point components obtained in the manufacture of tert-butyl methacrylate are, for example, tert-butyl acetate, tert-butanol and diisobutene.

To remove components with a low boiling point, the condensate from the catalyst removal step is fed to a distillation column for components with a low boiling point, in the lower part of which a liquid product is obtained from which components with a low boiling point have been removed, and in its upper part, vaporous components with a low temperature are obtained. boiling. The liquid product, from which the low boiling point components have been removed, containing predominantly the target ester and carboxylic acid, is fed for further processing. Condensation of vaporous components with a low boiling point is carried out. Preferably, the condensate is partially returned as a recycle stream to the top of the low boiling point distillation column.

The non-condensed, low boiling point vapor components may still contain up to 5% by weight of the desired ester based on the total overhead. The low boiling vaporous components also contain isobutene and are preferably fed to the esterification step.

The distillation temperature (liquid phase temperature) in the distillation column for low boiling point components is generally in the range of 30-110 ° C. The pressure is adjusted according to the product. In the case of tert-butyl methacrylate production, for example, the pressure is in the range of 0.1-0.25 bar (abs).

Suitable columns for the distillation of low boiling point components include conventional columns with irregular packing, or structured packing, or with bubble cap, valve or sieve trays. Preferably, however, a tray column with 30-50 double-flow trays is used. The low boiling point components are generally fed to the distillation column in the middle.

Preferably, the condensation of the low boiling point components is carried out in one or more condensers arranged in series, in particular in plate or shell-and-tube condensers. Preferably, shell and tube condensers are used with vertical tubes through which the steam flows from top to bottom.

Diisobutene is the major component of low boiling point removals. Diisobutene is a mixture of different isooctene isomers. In practice, these isomers have similar dew points. For example, the dew point of 2,4,4-trimethylpent-1-ene can be used as a guide.

In accordance with one embodiment of the invention, the condensation of the vaporous components with low boiling point occurs during two condensation stages, during which most of the diisobutene is condensed at a first temperature and the remainder of the diisobutene is condensed at a second temperature, which is lower than the first temperature. This provides the most efficient removal of diisobutene while minimizing co-condensation of unreacted isobutene. Fractional condensation preferably takes place in two series-connected capacitors.

After removing the acid catalyst (step b)) and removing low boiling point components (step c)) from the esterification mixture, the liquid containing tert-butyl ether is distilled in a still (step d)).

During purification by distillation in a still, the liquid containing tert-butyl ether is separated into a gaseous overhead containing tert-butyl ether and a liquid underflow containing carboxylic acid. At least partial, preferably complete, recirculation is carried out in the esterification step.

The distillation temperature is usually in the range of 40-130 ° C. The pressure is selected depending on the ether being distilled.

The distiller is usually a distillation column, such as a conventional tray column, such as a 30-50 double-flow tray column, and is fed in, for example, in the middle. Substantially pure target ether is removed through the top of the column. The process according to the invention can be carried out as continuous distillation, for example in a distillation column, or as batch distillation, for example in a distillation apparatus.

In the course of purification by distillation, at least in the upper part of the distiller, its walls are heated, and / or such walls are thermally insulated. As a result of heating or thermal insulation of the walls in the upper part, condensation of gaseous products on the walls is prevented; such condensation can cause the formation of unstable liquid phases prone to polymerization.

The configuration of the distiller usually includes a bell, which is connected through a flange to the cy- 6 035883 lindrical part of the distiller. The flange is a temperature bridge, therefore it can have a lower temperature than the rest of the upper part, therefore, there is an increased potential for condensation. Thus, further suppression of polymerization can preferably be carried out at this location.

The polymerization inhibitor is preferably fed into a flash tube through which a gaseous overhead product containing tert-butyl ether is withdrawn from the still. The supply of the polymerization inhibitor takes place in a metered manner, the metering is preferably carried out using a flow restrictor, which preferably provides a high efficiency of liquid distribution. The flow restrictors used are preferably a flow energy absorber, a valve, a converging pipe, a perforated plate, a nozzle, a capillary tube or a combination of these devices, in particular the flow restrictor is a nozzle. The use of a flow restrictor ensures high efficiency of distribution of the stabilizer at the walls in the upper part of the distiller, as well as at the flange connecting the bell and the cylindrical part of the distiller.

Due to the fact that the substances to be separated have different volatilities, the concentration distribution profile is different along the entire length of the column. The upper part of the distiller is dominated by tert-butyl ether, and the lower part is dominated by unreacted ethylenically unsaturated carboxylic acid. Polymerization inhibitors used for stabilization also have different volatilities. The boiling point of polymerization inhibitors is generally higher than the boiling point of tert-butyl ether. Therefore, evaporation and condensation of tert-butyl ether and / or carboxylic acid can occur at lower temperature locations in the still. In this case, condensate polymerization may occur, because the degree of suppression of polymerization is insufficient. The formation of these condensates is prevented by heating the walls at the top of the distiller and / or by thermally insulating these walls. Moreover, the even distribution of the polymerization inhibitor in the evaporator tube results in a more even distribution of the polymerization inhibitor on the walls in the upper part of the distillation unit, therefore, this leads to the formation of smaller amounts of condensate subject to polymerization.

The walls at the top of the still are preferably heated by thermal contact with a shell system through which a heating medium such as hot steam is passed.

The jacket system preferably includes an outer jacket, a semi-closed pipe, or hoses or heat tracing lines, in particular a semi-closed pipe. A semi-closed pipe is a semi-annular pipe segment that is located around the heated area and which is welded to the outer wall in this area. The semi-closed pipe preferably uses hot steam at a pressure of 1.2-2.5 bar, more preferably 1.3-1.7 bar, for example 1.5 bar. Usually the pressure is adjusted. Typically, the temperature is in the range of 105-130 ° C, preferably in the range of 108-113 ° C.

Another option is the use of electric heating, for example with heating wires or heating mats.

The wall temperature at the top of the still can also be kept above the dew point of the tert-butyl ether by thermal insulation. However, it is preferable to carry out active heating of the walls in the upper part of the still.

Thermal insulation of the walls at the top of the distiller can be provided using conventional insulation materials such as wool, glass wool, composite materials, panels, or sheets, or pipes that are attached to the outer wall at the top.

The temperature of the heated walls is preferably 2-20 ° C, more preferably 2-15 ° C, for example 5-10 ° C, higher than the dew point of the tert-butyl ether at the pressure prevailing in the still.

At least partial condensation of the gaseous overhead product containing tert-butyl ether occurs. In addition, at least a partial recirculation of the condensate to the distiller takes place as a return flow. The rest of the condensate is removed from the process as a product.

The higher the reflux ratio in the still, the more efficient the separation of ethylenically unsaturated acid and tert-butyl ether is. The term reflux ratio refers to the ratio between the amount of tert-butyl ether overhead condensate that is recycled to the distiller and the amount of tert-butyl ether overhead condensate that is fed to further steps. Accordingly, with a high reflux ratio, a low carboxylic acid content in the tert-butyl ether can be achieved. However, a high reflux ratio entails a high liquid phase temperature and a relatively high heat load on the ethylenically unsaturated carboxylic acid and tert-butyl ether. The reflux ratio is preferably in the range of 2-4, more preferably 2.5-3.5. The advantages of the method according to the invention under these conditions are manifested

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The total amount of acetic acid, propionic acid and ethylenically unsaturated carboxylic acid in the final product is preferably less than 300 ppm, more preferably less than 200 ppm, particularly preferably less than 130 ppm. It is particularly preferred if the total acid content in the final product is less than 300 ppm, more preferably less than 200 ppm, particularly preferably less than 130 ppm. To determine the acid content, the acid number is usually determined by titrimetric analysis, which is expressed as the methacrylic acid content.

Preferably, the condensation of the gaseous overhead product is carried out in one or more condensers arranged in series, in particular in plate or shell-and-tube condensers. Preferably, shell and tube condensers are used with vertical tubes through which the steam flows from top to bottom.

If a separate condenser is used, the pressure in the condenser is in the range 10-200 mbar abs., More preferably in the range 30-100 mbar abs., Particularly preferably in the range 50-90 mbar abs. Preferably the temperature in the condenser is 45-80 ° C, more preferably 50-65 ° C below the dew point of the tert-butyl ether at the pressure used. Preferably, the temperature is in the range of (-25) -0 ° C, more preferably in the range of (-20) to (-5) ° C.

The condensation temperature of tert-butyl acrylate at 60 mbar abs. Is, for example, 43 ° C. In this case, brine cooling can be used in the condenser.

In accordance with one embodiment of the invention, the condensation of the gaseous overhead product occurs in the form of a two-stage partial condensation. In this case, the most efficient removal of tert-butyl ether is ensured, while the co-condensation of the remaining components with a low boiling point is minimized. The two-stage partial condensation preferably takes place in two series-connected capacitors. The temperatures indicated here for partial condensation refer to the temperature of the condensate when it is removed from the respective condenser.

The temperature of the refrigerant in the second condenser is about 30-60 ° C lower than in the first condenser, where the temperature of the refrigerant is in the range of about 10-35 ° C.

The first temperature is preferably 0-45 ° C, more preferably 5-35 ° C below the condensation temperature of tert-butyl ether at the first pressure, and the second temperature is 45-80 ° C, preferably 50-65 ° C, particularly preferably 50 -55 ° C below the dew point of tert-butyl ether at the second pressure.

The second temperature is at least 5 ° C lower than the first temperature. The second temperature is preferably at least 10 ° C, more preferably at least 20 ° C, particularly preferably at least 30 ° C, most preferably at least 40 ° C below the first temperature.

Preferably, the first pressure is in the range 10-200 mbar abs., More preferably in the range 30-100 mbar abs., Particularly preferably in the range 50-90 mbar abs. Preferably, the second pressure is in the range 10-200 mbar abs, more preferably in the range 30-100 mbar abs, particularly preferably in the range 50-90 mbar abs. Usually, the first and second capacitors are connected on the gas side. In this case, the first and second pressures are the same.

The condensation temperature of tert-butyl acrylate at 60 mbar abs. Is, for example, 43 ° C. In this case, river water or cooling water brought to the required temperature can be used to cool the first condenser, while brine cooling can be used in the second condenser.

The (combined) condensate is partially removed from the process as a product.

The carboxylic acid-containing liquid underflow is at least partially recycled to the esterification step. The carboxylic acid-containing liquid underflow which is recycled to the esterification step is preferably cooled to a temperature in the range of 2540 ° C, preferably 25-35 ° C. This is usually done using a warm water cooler that operates at temperatures in the range 25-35 ° C, for example 30 ° C.

After cooling, the underflow substream, which is recycled to the esterification step, is preferably fed through a solids remover. Such a device is preferably a filter. In this way, you can ensure the removal of a large part of the undissolved solid impurities, such as, for example, polymer particles. This makes it possible, for example, to avoid clogging the lines. It is advisable to provide a filter with a bypass channel, the connection of which is carried out in such a way that the filter can be removed for cleaning or for replacement without stopping the process.

Another sub-stream of liquid underflow passes through a forced circulation evaporator, then it is withdrawn while passing through the heater, a superheated liquid recirculation stream is withdrawn from the heater, which is released into the distiller.

- 8 035883

The pressure at which the liquid underflow is withdrawn from the still and recirculated to the still is increased during superheating. The descent of the superheated recirculation stream is carried out through a flow restrictor. As a result, overheating occurs above the boiling point relative to the pressure inside the distiller.

As the superheated liquid passes through the flow restrictor and is fed back to the distiller, the liquid evaporates rapidly. This rapid evaporation, accompanied by a significant increase in volume, leads to an acceleration of the flow into the distiller, thereby reducing the risk of deposits, for example, polymers at the point of entry into the distiller. Therefore, it is preferable that the flow restrictor is located immediately upstream of the superheated liquid re-supply point to the distiller or even inside it.

In accordance with one embodiment of the invention, the operation of the distillation unit is carried out in a known manner with the metered supply of polymerization inhibitors (stabilizers) at appropriate points, for example in the upper part of the distillation unit. Thus, it prevents the ingress of vaporous compounds prone to polymerization or their condensates in tanks and pipelines that are not wetted with a stabilizer solution.

Preferably, when the forced circulation evaporator is operating, the flow rate of the liquid in the heating device, for example in the tube bundle of the heat exchanger, increases as compared to the flow rate during the operation of the natural circulation evaporator. This ensures that the surfaces of the heat exchanger are constantly cleaned by the flowing product. Another advantage of the invention is improved heat transfer between the heat exchanger and the heated fluid as a result of the increased flow rate, while this improved heat transfer in turn helps to prevent local overheating.

In carrying out the method according to the invention, a pump is used as a flow-moving device in order to provide forced circulation of the flow in the heating device. Such a pump is preferably positioned between the discharge line and the heater, thereby increasing the pressure of the flow to the heater to prevent evaporation of the liquid downstream of the heater. Another pump can be used to remove the underflow liquid stream.

Preferably, a decontamination composition is added to the stream which is fed to the flow transfer device, which is selected from, for example, succinimide derivatives commercially available under the name KOMAD to protect the device from contamination, clogging and consequent damage. In this way, it is possible to ensure that most of the undissolved solids, such as polymer particles, are removed before they enter the heater and cause fouling.

Conventional heat exchange devices can be used, preferably a heat exchanger in which indirect heat exchange with the heating medium takes place, in particular a shell and tube heat exchanger. Preferably, in the operating position, the heat exchanger is horizontal. Also, instead of shell and tube heat exchangers, plate or spiral heat exchangers can be used, since in this case, if the required flow rate is maintained and evaporation is suppressed, the same advantages can be provided as for a shell and tube heat exchanger.

It is also possible to heat the liquid discharged from the distiller using electrical energy.

The flow restrictors used preferably include a flow damper, a valve, a converting tube, a perforated plate, a nozzle, a capillary tube or a combination of these devices, in particular the flow restrictor is a valve. For example, a segmented butterfly valve can be used. It is particularly advantageous if it is possible to adjust the opening of the flow restrictor. By this, the pressure in the evaporator can be kept above the boiling pressure of the liquid relative to the pressure in the interior of the still in the event of a change in flow rate, which can occur, for example, during start-up and shutdown of the plant.

In a steady state, the reactants in the reactor are in the form of a solution in the target ether, which allows the reaction to be homogeneous and, particularly preferably, allows heat removal. Thus, for starting the reactor, it is preferably filled with the target ether. After that, reactants and catalyst are introduced into the reactor, and the reaction begins.

After starting up the plant, the contents of the reactor are preferably fed into a collecting vessel. The collecting tank is located at the lowest point of the installation, it is connected to the reactor through separate lines. Thus, it is possible to quickly empty the reactor in the event of a leak. As a rule, pumping systems are not required for this. The collection vessel is equipped with a pressure equalization device and is filled with an oxygen-containing gas with an oxygen content of 10 vol% or less, preferably 5 vol% oxygen or less in an inert gas, preferably nitrogen. The collecting tank is cooled using a pump and an external heat exchanger. The contents of the collection container can then be recycled separately.

- 9 035883

The reagents, in particular the carboxylic acid, are preferably used in substantially anhydrous form. The surfaces in contact with the reaction components in the process preferably consist of materials that are technologically resistant to the corrosive effects of the carboxylic acid used, for example stainless steel 1.4541 or 1.4571 or stainless steel grades having at least similar corrosion resistance characteristics. Due to the extremely low water content in the process system, even when strong inorganic acids are used as a catalyst, the corrosive effect remains within the range corresponding to the industrial resistance of these materials. In ethylenically unsaturated ester plants, cleaning typically needs to be carried out using a hot sodium hydroxide solution, exposing the materials used alternately with an organic acid and a sodium hydroxide cleaning agent. Therefore, in order to ensure the reliable operation of the device for a long time, it is preferable to use the so-called duplex steel, for example, steel grade 1.4462.

In those zones where the installation, in addition to the specified corrosive effects of inorganic acids and a strong inorganic acid used as a catalyst, is additionally subjected to high temperatures and mechanical stress, for example, in a thin-film evaporator to remove the acid catalyst from most organic substances, it is preferable to use materials with significantly better corrosion resistance, such as nickel-based materials such as 2.4602, 2.4605, 2.4610 or 2.4819. Experience has shown that these materials not only have a longer service life, but also a significant margin in the event of an unplanned appearance of water, which in this case acts as a factor contributing to corrosion, due to an even lower rate of removal of corrosive material compared to stainless steel. steel. The use of these materials provides advantageous characteristics in case of operation of the device in emergency conditions without the risk of a rapid arrival of the device in complete malfunction. In the case of operation of the device in an emergency situation, the presence of water can be caused, for example, by its accidental ingress into the system, for example, together with raw materials or auxiliary materials containing water, or as a result of loss of tightness in the reactor cooling system or in condensers used for fractional condensation. , or due to steam leakage into the process from devices that are heated directly using steam.

For cleaning the reactor, the empty reactor is preferably filled with a sodium hydroxide solution (eg 5% by weight in demineralized water) which is heated to about 80 ° C. and the solution circulated in the solution. This is followed by the disposal of cooled aqueous alkaline solutions, which remain after the completion of cleaning, if necessary, after appropriate treatment in a wastewater treatment plant (for example, in a water treatment plant). After the reactor has been cleaned, in particular when organic contaminants have been removed from it, residual solution in the reactor system or in plant components that require additional cleaning can be removed by flushing with water.

Isobutene is a highly flammable substance that, in the presence of oxygen, can form explosive mixtures that can ignite near hot surfaces at certain oxygen concentrations. Under standard operating conditions, the operation of the unit during start-up and shutdown is carried out in such a way that the oxygen concentration in the gas phase is constantly kept below the explosive limit. To this end, the installation is flushed and filled with an oxygen-containing gas with an oxygen content of 10 vol% or less, preferably 6 vol% oxygen or less, in a mixture with an inert gas, preferably nitrogen, before starting. Preferably, such an oxygen-containing gas is oxygen-poor air with an oxygen content of 10 vol% or less, which is obtained, for example, by diluting the air with molecular nitrogen. Preferably, all components are fed into the process in an oxygen-poor atmosphere. The complete elimination of oxygen is undesirable, particularly when oxygen is required to ensure the effectiveness of one of the stabilizers described below. If oxygen is consumed during the process, preferably at appropriate points, for example at the bottom of the still, a continuous supply of fresh oxygen-poor air is carried out. When using oxygen-poor air, the composition of the gas phase does not reach the explosive range of oxygen concentration even in the case of non-uniform composition of the gas phase.

In order to detect air leaks, in particular in plant components operating at reduced pressure, instruments for on-line oxygen measurement are preferably installed at various points in the plant. More preferably, such instruments for on-line oxygen measurement are installed in the non-condensable vapor lines obtained from the fractional condensation step.

The reactor is completely filled with liquid, therefore, it should preferably be protected against thermal expansion in the form of a safety valve. In addition, the reactor should preferably be provided with a system for rapid isolation, emptying and de

- 10 035883 compression (SAEES), with the help of which the entire contents of the reactor can be dumped in the event of loss of tightness without contact with the environment into a ventilated collection vessel, the safe evacuation of which can be carried out in the event of an explosion hazard. The cooling of the contents of such a collecting vessel can preferably be carried out by means of a heat exchange device in order to enable controlled removal, if necessary, of the heat generated by the further reaction. The collecting container and its special devices are preferably configured so that its contents can be fed back into the process at various points.

The carboxylic acids used in the present process, if they have ethylenically unsaturated groups, can be substantially polymerizable, in particular at relatively high temperatures. These compounds are often exposed to high temperatures, in particular during the distillation process, which can lead to undesirable free radical polymerization. This polymerization results in fouling of the device, clogging of pipelines and pumps, and accumulation of deposits on the column plates and the surfaces of the heat exchanger. The cleaning of the installation requires a lot of effort and expense, and can also cause environmental pollution, in addition, due to the need for cleaning, the power utilization of the installation decreases. Secondly, due to uncontrolled free radical polymerization, there can be a threat to industrial safety. This type of polymerization can be prevented by using appropriate stabilizers.

Polymerization stabilizers or inhibitors are generally used to suppress polymerization. Such stabilizers are usually solids that are fed into the process as a solution. The stabilizer solutions are preferably prepared in batches.

Suitable stabilizers are, for example, N-oxyl compounds, nitroso compounds, phenolic compounds, phenothiazines and mixtures thereof. The polymerization inhibiting activity of stabilizers is generally enhanced in the presence of molecular oxygen. In some cases, the presence of molecular oxygen is absolutely necessary for the stabilizer to be effective. Therefore, it is preferred that molecular oxygen is present in the plant.

Suitable N-oxyl compounds include 1-oxyl-2,2,6,6-tetramethylpiperidine (TEMPO), 1-oxyl-2,2,6,6-tetramethylpiperidin-4-ol (4-HT), 1-oxyl- 2,2,6,6-tetramethylpiperidin-4-one, 1-oxyl2,2,6,6-tetramethyl-4-n-propoxypiperidine, 1-oxyl-2,2,6,6-tetramethyl-4- (2 -methoxyethoxy) piperidine, 1-oxyl-2,2,6,6-tetramethyl-4- (2-methoxyethoxyacetoxy) piperidine; 1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl stearate, 1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl acetate, 1-oxyl-2,2,6,6tetramethylpiperidine- 4-yl butyrate, 1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl 2-ethylhexanoate, 1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl octanoate, 1-oxyl-2,2 , 6,6-tetramethylpiperidin-4-yl laurate, 1oxyl-2,2,6,6-tetramethylpiperidin-4-yl benzoate, 1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl 4-tert-butyl benzoate, bis- (1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl) succinate, bis- (1-oxyl-2,2,6,6tetramethylpiperidin-4-yl) adipate, bis- (1-oxyl -2,2,6,6-tetramethylpiperidin-4-yl) sebacate, bis- (1oxyl-2,2,6,6-tetramethylpiperidin-4-yl) n-butylmalonate, bis- (1-oxyl-2,2 , 6,6-tetramethylpiperidin-4yl) phthalate, bis- (1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl) isophthalate, bis- (1-oxyl-2,2,6,6-tetramethylpiperidine -4-yl) terephthalate, bis- (1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl) hexahydroterephthalate, 1-oxyl-2,2,6,6-tetramethyl-4-allyloxypiperidine, 1-oxyl -2,2,6,6-tetramethyl-4-acetamidopiperidine, 1oxyl-2,2,6,6-tetram methyl-4- (N-butylformamido) piperidine, ^^ - bis- (1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl) adipamide, N- (1-oxyl-2,2,6,6 -tetramethylpiperidin-4-yl) caprolactam, N- (1-oxyl2,2,6,6-tetramethylpiperidin-4-yl) dodecyl succinimide, 2,4,6-tris- (N-butyl-N- (1-oxyl-2 , 2,6,6 tetramethylpiperidin-4-yl] -s-triazine, 4,4'-ethylenebis- (1-oxyl-2,2,6,6-tetramethylpiperazin-3-one), 1 oxyl-2,2, 6,6-tetramethyl-4- (2,3-dihydroxypropoxy) piperidine, 1-oxyl-2,2,6,6-tetramethyl-4- (2hydroxy-4-oxapentoxy) piperidine, di-tert-butylnitroxyl and 4, 4 ', 4-tris- (2,2,6,6-tetramethylpiperidinooxyl) phosphite.

1-oxyl-2,2,6,6-tetramethylpiperidin-4-ol (4-HT) can be used particularly effectively.

Suitable nitroso compounds include nitrosophenol, N-nitrosodiphenylamine, isoamyl nitrite, N-nitrosocyclohexylhydroxylamine, N-nitroso-N-phenylhydroxylamine, and salts thereof.

Suitable phenolic compounds include hydroquinone, 2-methoxyphenol, 3-methoxyphenol, 4-methoxyphenol (MEHQ), 2-ethoxyphenol, 3-ethoxyphenol, and 4-ethoxyphenol. 4-Methoxyphenol (MEHQ) can be used especially effectively.

Suitable phenothiazines include phenothiazine (PTZ), 2-methylphenothiazine, 2-octylphenothiazine, 2-nonylphenothiazine, 2,8-dimethylphenothiazine, 3,7-dimethylphenothiazine, 3,7-diethylphenothiazine, 3,7-dibiazinyl-3,8 dioctylphenothiazine, 3,7dinonylphenothiazine, 2,8-dinonylphenothiazine, 2- (a, a-dimethylbenzyl) phenothiazine, 3,7-bis- (a, adimethylbenzyl) phenothiazine and 2,8-bis- (a, a-dimethylbenzyl) phenothiazine ... Phenothiazine (PTZ) can be used especially effectively.

It is also possible to use several stabilizers. Stabilizers are usually used

- 11 035883 in an amount of about 2-2000 ppm based on the total amount of carboxylic acid and isobutene.

Preferably, the stabilizer is added in solution with a solvent. Suitable solvents are, in principle, all solvents in which a particular stabilizer can dissolve and which can be miscible with the liquid phase to be stabilized. The solvent can be a liquid that is in any case used in the process, primarily to avoid contamination by (external) solvents that are not needed for the process, or so that there is no need to remove external solvent. Especially preferably, pure target ester is used as solvent.

When the stabilizer is supplied, as a rule, its amount is regulated by means of pumps; preferably, for better distribution, the stabilizer solution is sprayed with spray devices such as spray nozzles.

Some of these stabilizers are effective only in the presence of oxygen, for example MEHQ, so it is preferable, for example, if the air contains a relatively high concentration of oxygen. On the other hand, the oxygen concentration must be limited to relatively low values to avoid the formation of explosive mixtures. The process is carried out in such a way that the concentration of oxygen in the gas phase at all relevant points is constantly kept below the explosive limit. The volume ratio of oxygen to nitrogen in all gas mixtures used in steps a) to d) is preferably in the range of 0.03-0.11.

In particular, the liquid phase tends to polymerize at lower concentrations of stabilizers, and in some cases at lower oxygen concentrations. Since stabilizers are generally not volatile, they accumulate at the bottom of a particular evaporation system during the evaporation stages. Consequently, it is usually necessary to add the stabilizer after evaporation of the polymerizable compounds when condensation of the compounds occurs, since the resulting condensate usually contains practically no stabilizer.

The method according to the invention includes many technological stages, which use different mixtures of substances with different compositions under different technological conditions. In order to ensure the safety and economic feasibility of the plant, it is necessary to vary the stabilizers added in each case, which are fed into the process at different points.

In a preferred embodiment of the invention, in the case of the reaction of the carboxylic acid with isobutene, a stabilizer selected from phenothiazines is used, more preferably PTZ is used. The (meth) acrylic acid used may already have been pre-stabilized by PTZ, which is especially preferred when starting up the installation. Additional quantities of PTZ can be metered into the reactor. During the partial evaporation of the esterification mixture to remove the acid catalyst, the PTZ is distilled into a liquid phase of the high boiling point components, which is separated from the main stream containing the product. The liquid phase of the high boiling point components is preferably recirculated back to the reactor, thus, as a rule, only small additional amounts of fresh PTZ need to be fed continuously to the process.

In a preferred embodiment of the invention, a stabilizer selected from N-oxyl compounds is added to the condensation step. More preferably, the 4-HT solution is added to the target ester. The stabilizer is preferably fed at the point where steam is fed to the condenser, in the case of a two-stage partial condensation at the point where the steam is fed to the first condenser. This stabilizer is also fed to the second condenser along with the condensate recycle stream at the point where steam is fed to the second condenser.

Preferably, the stabilization of the low boiling point distillation column is carried out using a stabilizer selected from N-oxyl compounds, more preferably using a stabilizer that is 4-HT. The initial stream supplied to the column for distillation of components with a low boiling point contains 4-HT from the previous stage, an additional amount of 4-HT is added at the top of the column for distillation of components with a low boiling point, in particular at the point of supply of the first condenser, and passes into the column distillation of components with a low boiling point together with a return flow of condensate.

In a preferred embodiment, the stabilizer is selected from Noxyl compounds, more preferably 4-HT, the stabilizer being fed into the inlet of the still.

Stabilization of the bottom and stripping section of the distillation unit takes place using N-oxyl compounds. Preferably, the stabilization of the distillation section of the still is not carried out with N-oxyl compounds, since they cannot be completely prevented from entering the product. The presence of N-oxyl compounds in the target ester is undesirable, as they can lead to a loss of color in the product and substances derived from it. Thus, a stabilizer selected from phenol compounds is fed to the rectification section, in particular, the stabilizer is

- 12 035883 is MEHQ. This stabilizer is also used to stabilize the product, therefore it has no negative effect and does not need to be removed in subsequent steps. Preferably, MEHQ is added to the recycle stream through the condensers and / or is added to the condensate return stream to the distillation unit. This is done suitably by injection through a nozzle located at the center of the outlet of the evaporator tube.

In order to ensure the efficiency of the MEHQ, a gas containing molecular oxygen, preferably oxygen-poor air (5 vol% oxygen in nitrogen), is preferably fed to the bottom of the still. These measures prevent the formation of polymers in the condensers, evaporator tubes and in the distiller, or at least prevent it to such an extent that it is possible to economically operate the installation without the need to temporarily shut down for cleaning purposes.

The present invention is illustrated in more detail by the following examples and diagrams.

FIG. 1 shows a schematic diagram of an installation for carrying out the method according to the invention.

FIG. 2 shows a schematic diagram of a distiller for distillation purification.

FIG. 3 shows the differential pressure in the still according to FIG. 2, which uses a natural circulation evaporator without heating the top of the still, as a function of time.

FIG. 4 shows the heating power in the distiller according to FIG. 2, which uses a natural circulation evaporator to heat the top of the still, in the form of steam consumption as a function of time.

FIG. 5 shows the differential pressure and heating power (in terms of steam consumption) in the distiller in accordance with FIG. 2, which uses a forced circulation evaporator with heated top of the distiller, functionally versus time.

Referring to FIG. 1, the mixing device M1 supplies the aliphatic C1-C4 carboxylic acid, the stabilizer I1 and the acid catalyst in the form of a mixture to the reactor R1 through line 1 and the nozzle E1 (not shown in FIG. 1). Isobutene is fed to the bottom of the reactor R1. Through the nozzle E1, non-condensed vapor streams B3 and B6 containing isobutene were fed into the reactor R1 from condensers C2 and C4. The condensate from the reverse flow condenser W1 is fed to the reactor R1.

In the reactor R1, the reaction of addition of isobutene and aliphatic C1-C4 carboxylic acid takes place. The reactor has four cooled reaction zones. The reaction zones are separated by means of dividing plates, the transition from one reaction zone to another is a small hole. The reagents were mixed in the reactor using the E1 nozzle and as a result of vortices during the transition from one zone to the next.

The liquid reaction product G1 is withdrawn at the top of the reactor R1 by means of a level controller to create a constant liquid-gas interface. The gas phase, consisting mainly of inert gases, isobutene and minor amounts of tert-butyl ether, is fed to the reverse flow condenser W1 through line 2. The condensate from the reverse flow condenser W1 contains isobutene and acrylic acid and is fed to the reactor R1 through line 3. Gas phase from the condenser W1, the reverse flow is removed from the process as an off-gas via line 4.

The liquid reaction product G1 is withdrawn through the side in the upper part of the reactor R1 and is fed with controlled quantity to an evaporation unit V1, consisting of a falling film evaporator and a separation vessel (not shown separately in FIG. 1). The pressure of the liquid reaction product is reduced by a throttle valve (not shown in Fig. 1) from the pressure in the reactor to a reduced pressure, at which the catalyst is then removed. In the falling film evaporator of the evaporation unit V1, the reaction mixture is partially evaporated and fed further to the separation vessel. The separation vessel preferably includes a droplet separator to sufficiently reliably remove high boiling point components such as sulfuric acid and stabilizer I1 entrained in the reaction mixture. The non-gaseous components are collected in the separation vessel as the first phase of the high boiling point SPh1 components and cooled with an external cooler (not shown in FIG. 1) in order to prevent the present tert-butyl ester from re-reacting to produce carboxylic acid and isobutene.

A portion of the first phase of the components with high boiling point SPh1 is fed, while adjusting the amount, to the thin-film evaporator V2 in order to further remove the carboxylic acid or tert-butyl ether in gaseous form. The gas phase obtained in the thin-film evaporator V2 is recirculated to the separation vessel of the evaporation unit V1 via line 5, while part of the second liquid phase is fed to the settler A1 by the high boiling point component SPh2. Preferably, substreams of the second phase of the components with a high boiling point SPh2 are used to preheat the feed stream fed to the thin film evaporator of the evaporation unit V1. By varying the hot substreams, it is possible to vary the composition of the feed stream supplied to the thin film evaporator of the evaporation unit V1 and the temperature of the feed stream.

- 13 035883

Another part of the first phase of the components with a high boiling point SPh1 and another part of the second phase of the components with a high boiling point SPh2 are recycled to the reactor R1 jointly or separately in each case, with the regulation of the amount of phase, through the nozzle E1 (recirculation of the second phase of the components with high boiling point SPh2 is not shown in Fig. 1).

In condensers C1 and C2, fractional condensation of gaseous components from the separation tank of the evaporation unit V1 occurs, steam B1 from condenser C1 is fed to condenser C2: The addition of stabilizer I2 is carried out in the upper part of condenser C1, and stabilizer I3 is added in the upper part of condenser C2. To cool the condenser C1, for example, river water or cooling water with the same temperature can be used, while in the condenser C2, brine cooling is used. Steam B3, which has not condensed in the condenser C2, is fed to the reactor R1 through the nozzle E1.

The condensates K1 and K2 obtained in the condensers C1 and C2 are combined and fed to the side of the distillation column D1. Distillation column D1 removes low boiling point components, in particular diisobutene and isobutene. The bottom of the distillation column D1 is heated by a circulating evaporator (not shown in FIG. 1), thereby pumping part of the circulating underflow. The low boiling point components B4 are removed as vapor at the top of the distillation column D1 and are fractionally condensed in condensers C3 and C4. Steam B5 from condenser C3 is fed to condenser C4. To cool condenser C3, for example, river water or cooling water at the same temperature can be used, while condenser C4 uses brine cooling. Steam B6, which has not condensed in the condenser C4, is fed to the reactor R1 through the nozzle E1. The addition of the stabilizer I4 is carried out at the top of the capacitor C3. Condensates K3 and K4 obtained in capacitors C3 and C4 are combined; one substream is fed to the distillation column D1 as a reflux stream and the remainder is fed to the settler A2.

The bottom stream S1 from the distillation column D1 is fed to the side of the distillation column D2. The stabilizer I5 is dosed into the inlet of the distillation column D2. The bottom of the distillation column D2 is heated by a circulating evaporator (not shown in FIG. 1) through which a portion of the circulating underflow is pumped. During forced circulation, oxygen-poor air is supplied to the bottom of the distillation column D2 by means of a pump.

In the distillation column D2, the tert-butyl ether is separated from the remaining aliphatic carboxylic acid. Typically, the boiling point of the carboxylic acid is higher than the boiling point of the tert-butyl ether, so pure tert-butyl ether is withdrawn from the top, and the carboxylic acid is obtained at the bottom of the distillation column D2. In order to avoid condensation of tert-butyl ether at the top of the column, the top of the column was heated with steam. This also prevents the polymerization of the tert-butyl ether, which in some cases can occur due to condensation. The bottoms stream S2 from the distillation column D2 is recycled to the reactor R1 through a heat exchanger (not shown in FIG. 1) and optionally through a filter (not shown in FIG. 1). A bypass channel can be provided for the filter so that the filter can be removed for cleaning or replacement without stopping the process.

In the condensers C5 and C6, the fractional condensation of steam B7 from the distillation column D2 occurs, the steam B8 from the condenser C5 is fed to the condenser C6. The addition of the I7 stabilizer is carried out at the top of the C5 capacitor, and the I8 stabilizer at the top of the C6 capacitor. Steam B9, which has not condensed in condenser C6, is removed from the process as off-gas. The off-gas can, for example, be discharged into a flare or into an off-gas incinerator.

The sub-stream of the combined condensates P1 and P2 from condensers C5 and C6 is fed to condensers C5 and C6 (not shown in FIG. 1) or, with the addition of stabilizer I6, is fed as a reflux stream to distillation column D2. Another substream of combined condensates P1 and P2 from condensers C5 and C6 is withdrawn from the process as pure tert-butyl ether through a heat exchanger (not shown in FIG. 1). Additional stabilizer I9 can be added to pure tert-butyl ether to stabilize during storage.

The reactor is preferably provided with a fast isolation, emptying and decompression (SAEES) system, by means of which the entire contents of the reactor R1 can be discharged in the event of loss of tightness into a vented collecting vessel (not shown in FIG. 1). Cooling of the contents of such a collecting vessel can be carried out using a heat exchange device in order to remove the heat generated by the further reaction. The contents of the collection vessel can be fed back into the process at various points, in particular to the reactor R1, to the falling-film evaporator V2 or to the thin-film evaporator of the evaporation unit V1.

FIG. 2 depicts a preferred embodiment of the invention shown in FIG. 1. Referring to FIG. 2, the bottom stream S1 from the distillation column D1 is fed to the side of the distillation column D2. Dosed addition of I5 stabilizer to the inlet

- 14 035883 distillation column D2. The sub-stream of the combined condensates P1 and P2 from the condensers C5 and C6 (not shown in Fig. 2) with the addition of the stabilizer I6 is fed as a reflux to the distillation column D2.

The bottom of the distillation column D2 is heated by a forced circulation evaporator Z1, through which a portion of the underflow S2 is pumped by circulation by a pump V1. The heated stream from the recycle evaporator is recirculated through nozzle E2 to distillation column D2. Oxygen-poor air L1 is introduced to the bottom of the distillation column D2.

In the distillation column D2, the tert-butyl ether is separated from the remaining aliphatic carboxylic acid. Typically, the boiling point of the carboxylic acid is higher than the boiling point of the tert-butyl ether, so pure tert-butyl ether is withdrawn from the top, and the carboxylic acid is obtained at the bottom of the distillation column D2. In order to avoid condensation of tert-butyl ether at the top of the column, the top of the column was heated with steam. This also prevents the polymerization of the tert-butyl ether, which in some cases may occur due to condensation. This polymerization is also prevented by the metered addition of inhibitor I10 to evaporator tube B7 using nozzle E3.

Example 1.

The example below was carried out in a plant according to FIG. 1. Except where otherwise indicated, all percentages are by weight. Conversion of methacrylic acid and isobutene to tert-butyl methacrylate was carried out with the addition of sulfuric acid. The installation was flushed and filled with oxygen-poor air (5 vol.% Oxygen in nitrogen).

Methacrylic acid (MAA, 99.97%, 0.012% acetic acid and 0.011% propionic acid, 575 kg / h) was mixed in mixer M1 with methacrylic acid recirculated from the bottom of distillation column D2 (91.35% MAA, 5.15 % of components with high boiling point, 3.32% tert-butyl methacrylate (TBMA), 485 kg / h) and with a solution of phenothiazine (PTZ) (97,97% MAA, 2,10% PTZ, 0,01 m ³ / h ) as a stabilizer for I1. Sulfuric acid (technical grade 96%, 3.7 kg / h) was dosed using nitrogen (3.5 bar abs). In addition, a substream (1500 kg / h) of the liquid phase was fed into the mixture from the separation vessel of the evaporation unit V1.

Liquid reactants and return streams were fed as a mixture through line 1, a cooler (shell and tube system, 180 m ² , 1.4571 stainless steel, not shown in FIG. 1) and nozzle E1 to reactor R1, a cascade bubble column. The outlet temperature from the cooler was 33 ° C.

Isobutene (402 kg / h) was dosed directly into the bottom of the reactor. In addition, the vaporous streams B3 and B6 containing isobutene from condensers C2 and C4 (122 m ³ / h) were dosed into the reactor R1 through the nozzle E1 from condensers C2 and C4 (122 m ³ / h), together with the condensate containing tert-butyl methacrylate from the reverse flow condenser W1. The E1 nozzle was a jet nozzle. The pressure in the nozzle was raised to approx. 2 bar abs by the moving jet.

Reactor R1 had four reaction zones, the transition between the zones was provided by means of a small section hole (diameter 24 mm). Each of the zones was cooled (zone 1: external cooler using river water, 121 kW; zone 2: external brine cooler, 28 kW; zone 3: internal brine cooler, 14 kW; zone 4: internal brine cooler, 14 kW ; the temperature of each of the brines was -20 ° C). In reactor R1, the reaction of isobutene and methacrylic acid addition took place with the release of a small amount of heat (-37.6 kJ / mol) at a temperature of 35 ° C in zone 1, 25 ° C in zone 2, 20 ° C in zone 3, and 16 ° C in zone 4 and at 1.92 bar abs.

Mixing of reagents took place in the reactor, initially with the help of nozzle E1, and then as a result of vortices during the transition from one zone to the next. In the upper part of the column (zone 4), the boundary between the liquid and gas phases was established by means of level control in a closed loop.

The gas phase contained 41.18% isobutene, 0.74% tert-butyl methacrylate (TBMA) and inert gases, and was fed to the reverse flow condenser W1 through line 2. Condensation of TBMA entrained in the exhaust gas was carried out using the reverse flow condenser W1 and TBMA was recirculated to reactor R1 via line 3 mixed with isobutene-containing vapors from condensers C2 and C4 via nozzle E1. The pressure at the top of the reactor R1 in the off-gas line was adjusted to 1.21 bar abs. A pressure of 1.92 bar was set at the bottom of the reactor. Gaseous components from the reverse flow condenser W1 (4.37 m ³ / h) were withdrawn from the process via line 4.

The liquid reaction product from reactor R1 had the following composition:

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8.58% isobutene

41.63% MAA

44.00% TBMA 0.68% diisobutene 2.66% high boiling components 2.10% sulfuric acid 0.35% other components

The liquid reaction product from the reactor R1 (3.39 m ³ / h, 16 ° C) was removed at the top of the reactor R1 and fed through a basket filter (0.1 m ² , not shown in Fig. 1) into the evaporator with a falling foil (71.5 ° C, 309 kW, 47 m ² ) of the V1 evaporation unit. The pressure was reduced to 60 mbar abs using a closed loop control valve (flow control). As a result of the evaporation of some of the components with a low boiling point, a two-phase mixture was formed. In the falling film evaporator of the evaporation unit V1, the two-phase mixture was further evaporated with temperature control at 55.2 ° C and 70 mbar abs., And it was fed further into the separation vessel of the evaporation unit V1. The separation vessel was equipped with a droplet eliminator for sufficiently reliable removal of sulfuric acid and PTZ.

The non-gaseous components in the separation vessel of the evaporation unit V1 were cooled to 2 ° C. with a brine cooler using a pumped circulation stream as the first phase of the high boiling point components SPh1. In the separation vessel, the mixing temperature was suitably maintained at approximately 8-10 ° C. A portion of the pumped circulation stream (1500 kg / h) of the first phase of the high boiling point components SPh1 was fed back to the reactor R1 to recycle the sulfuric acid. In addition, a part of the pumped circulation of the first phase of the components with a high boiling point SPh1 (82 kg / h) was fed to a thin-film evaporator V2 (4 m ² , chromium-nickel-molybdenum alloy 2.4610) to remove additional amounts of valuable product (TBMA, MAA) through the upper part of thin film evaporator V2 (89 ° C, 70 mbar abs). The thin film evaporator V2 was heated with low pressure steam. Downstream of the underflow outlet of the thin film evaporator V2, a pump (not shown in FIG. 1) was connected to feed the second phase sub-stream of the high boiling point components SPh2 for discharge to the settler A1. On the way to the settling tank A1, the second phase substream of the high boiling point components SPh2 was cooled from 89 to 35 ° C. This was done using a sheathed pipe through which water at a temperature of 30 ° C was fed in the opposite direction.

Another substream of the second phase of the high boiling point components SPh2 from the thin-film evaporator of the evaporation unit V2, in turn, was mixed in the form of a hot recycle stream directly with the original stream entering the thin-film evaporator of the evaporation unit V2. By varying the recirculation flow rate, it was possible to control the flow rate and temperature of the feed stream over a wide range. Combined with the regulation of the amount and temperature of the heating steam, the V2 thin film evaporator was able to withstand a wide range of loads.

Another second phase sub-stream of high boiling point components SPh2 from the thin film evaporator was added to the cold pumped circulation from the pump suction side to the thin film evaporator V2, but this resulted in only slight heating of the pumped circulation stream. The initial flow to the thin film evaporator V2 was withdrawn from the discharge side of the pump.

Steam from the thin film evaporator V2 was fed through line 5 to the separation vessel of the evaporation unit V1. Steam B1 from the separation vessel of the evaporation unit V1 (at a temperature of approximately 68 ° C) had the following composition:

57% TVMA

24% MAA

16% isobutene

3% other components

Fractional condensation of steam B1 was carried out; for this purpose, it was fed to the upper part of condenser C1 (shell-and-tube heat exchanger, 75 m ² , cooling: river water (27 ° C), 60 mbar abs., Stainless steel 1.4571). In condenser C1, the feed mixture was cooled to 29 ° C.

Steam B2 from condenser C1 (containing approximately 35% TBMA, 5% MAA, 60% isobutene) was fed to the top of condenser C2 (shell and tube heat exchanger, 30 m ² , cooling: cooling brine (-20 ° C), 60 mbar abs., stainless steel 1.4571). Condensate K2 from capacitor C2 (containing approximately 86% TBMA, 5% MAA, 4% isobutene, approximately 17 ° C) was combined in containers (not shown in FIG. 1) with condensate K1 from capacitor C1. Steam B3 from condenser C2 (co- 16 035883 containing approximately 95% isobutene) was mixed using a pump (not shown in Fig. 1) with steam

B6 from condenser C4 and recycled to reactor R1.

Condensate K1 from capacitor C1 (containing approximately 68% TBMA, 28% MAA, 0.7% isobutene) was combined in a vessel (not shown in FIG. 1) with condensate K2 from capacitor C2. The combined condensate from C1 and C2 had the following composition:

68.09% TBMA

28.17% MAA

1.13% diisobutene

0.74% isobutene

0.61% other components

The sub-stream of combined condensates from C1 and C2 was fed together with a solution of 4-hydroxy-TEMPO (4-HT) (2% in TBMA) as a stabilizer I2 to the top of the condenser C1; in turn, its substream was fed as a stabilizer I3 to the top of the capacitor C2.

Another substream of the combined condensates from condensers C1 and C2 was fed to distillation column D1 (40 double-flow trays, 91 ° C at the bottom of the column, 120 mbar abs. In the overhead part of the column) to tray 23. The distillation column D1 was heated using an evaporator with natural circulation (4 bar abs., steam). The temperature of the distillation column D1 was controlled by a control valve in the return line. The vacuum pressure was controlled by a control valve in the suction line to the vacuum unit.

Fractional condensation of steam B4 from distillation column D1 was carried out; for this purpose, it was fed to condenser C3 (shell-and-tube heat exchanger, 110 m ² , cooling: river water (27 ° C), 120 mbar abs., Stainless steel 1.4571). In condenser C3, the feed mixture was cooled to 29 ° C. Condensate K3 from capacitor C3 was combined in tanks with condensate K4 from capacitor C4.

Steam B5 from condenser C3 was fed to condenser C4 (shell and tube heat exchanger, 8 m ² , cooling: cooling brine (-20 ° C), 120 mbar abs., Stainless steel 1.4571) and cooled to -2 ° C. Condensate K4 from capacitor C4 was combined in a container (not shown in Fig. 1) with condensate K3 from capacitor C3. B6 pairs of capacitor C4 (65,12 m ³ / h of isobutene 62.94%) were mixed by a pump (not shown in FIG. 1) to B3 of the capacitor C2 vapor and condensate from condenser W1 Reverse flow and recycled to the reactor R1 ...

A sub-stream of the combined condensates C3 and C4 was fed to the top of the distillation column D1; in turn, its substream was fed as a mixture with a 4-HT solution (2% in TBMA) as stabilizer I4 to the top of condenser C3.

The bottom product from the distillation column 1 had the following composition:

69.63% TBMA

28.65% MAA

1.72% other components

Bottom product S1 from distillation column D1 was mixed with 4-HT solution (2% in TBMA) as stabilizer I5 and fed to distillation column D2 (40 double-flow trays, 99 ° C column bottom, 60 mbar abs. At column head ) into tray 18. Distillation column D2 was heated using a circulating evaporator (4 bar abs., steam). The evaporator used natural or forced circulation (see options 1-3). The temperature of the distillation column D2 was controlled by a control valve in the return line. The vacuum pressure was controlled by a control valve in the suction line to the vacuum unit.

At the bottom of the distillation column D2, a metered addition of oxygen-poor air of 6 m ³ / h was performed. (5 vol.% Oxygen in nitrogen).

Fractional condensation of steam B7 from distillation column D2 (containing 99.83% TBMA) was carried out, for this purpose it was fed to condenser C5 (shell-and-tube heat exchanger, 72 m ² , cooling: river water (27 ° C), 60 mbar abs., Stainless steel grade 1.4571). In condenser C5, the feed mixture was cooled to 29 ° C. The condensate P1 from the capacitor C5 was combined in a container (not shown in Fig. 1) with the condensate P2 from the capacitor C6.

Steam B8 from condenser C5 was fed to the top of condenser C6 (shell and tube heat exchanger, 12 m ² , cooling: cooling brine (-20 ° C), 55 mbar abs., Stainless steel 1.4571) and cooled to -17 ° C. Condensate P2 from condenser C6 was combined in tanks (not shown in FIG. 1) with condensate P1 from condenser C5 as product. Steam B9 from condenser C6 was withdrawn from the process using a pump (not shown in FIG. 1).

The sub-stream of the combined condensates P1 and P2 from condensers C5 and C6 was fed to distillation column D2 as a reflux with the addition of a solution of 4-methoxyphenol (MEHQ, 2% in TBMA) as stabilizer I6. Other substreams of combined condensates P1 and P2 from capacitors C5 and C6 were fed to capacitors C5 and C6, respectively, with the addition of a solution of 4methoxyphenol (MEHQ, 2% in TBMA) as stabilizers I7 and I8.

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In this case, the column was stabilized using a high content

MEHQ, while the 4-methoxyphenol content of capacitors C5 and C6 was 15 +/- 5 ppm.

In order to avoid condensation of TBMA at the top of the distillation column D2, which can also lead to TBMA polymerization, the top of the column was heated with steam (4 bar abs).

Another substream of combined condensates P1 and P2 from condensers C5 and C6 after pressure increase to 4 bar abs. cooled to 20 ° C using a heat exchanger (spiral heat exchanger, cooling: cooling brine (-20 ° C), not shown in Fig. 1) and withdrawn from the process as a product. Its substream was used as a solvent for the 4-HT and MEHQ stabilizers.

The product had the following composition:

99.68% TBMA

0.30% isobutene

100 ppm acid (MAA, acetic acid, propionic acid)

ppm MEHQ

Bottom S2 from distillation column D2 (containing 91.35% MAA) after pressurizing to 4 bar abs. cooled to 35 ° C using a heat exchanger (spiral heat exchanger, 5 m ² , cooling: warm water, stainless steel 1.4571, not shown in Fig. 1), and the substream was combined with the original stream of methacrylic acid and fed to the reactor R1.

The reduced pressure required in the evaporation unit V1 and downstream units was created using a vacuum unit. Used gear piston compressors without lubricating oil.

To obtain a stabilizer, a solution of phenothiazine, methacrylic acid in pure form was first fed into a mixing vessel (heated in a tube with water, 30 ° C, ventilated). The PTZ was fed in solid form into a vessel in which mixing took place, through an emptying station and through a pneumatic system for transferring finely dispersed products using a pressure reduction using a pump. Small amounts of oxygen-poor air (5 vol.% Oxygen in nitrogen) were added. With stirring, the PTZ was dissolved, and the PTZ solution was fed into a reservoir tank (heated in a tube with water, 30 ° C, vented), from which it was dosed into the process.

To obtain the stabilizer, a solution of 4-hydroxy-2,2,6,6-tetramethyl-1-oxyl-piperidine, tert-butyl methacrylate (from the combined condensates P1 and P2 from condensers C5 and C6) was first fed into a mixing vessel (vented). The supply of 4-HT was carried out in solid form into a vessel in which mixing took place through an emptying station and through a pneumatic system for transferring fines by means of pressure reduction using a pump. Small amounts of oxygen-poor air (5 vol.% Oxygen in nitrogen) were added. With stirring, the 4-HT was dissolved, and the 4-HT solution was fed into a reservoir vessel (ventilated) from which it was dosed into the process.

To obtain the stabilizer, a solution of 4-methoxyphenol, tert-butyl methacrylate (from the combined condensates P1 and P2 from condensers C5 and C6) was first fed into a mixing vessel (ventilated). MEHQ was fed in a solid form into a container, in which mixing took place, through an emptying station and through a pneumatic system for transferring fine products using a pressure reduction using a pump. Small amounts of oxygen-poor air (5 vol.% Oxygen in nitrogen) were added. With stirring, the MEHQ was dissolved, and the MEHQ solution was fed into a reservoir vessel (ventilated), from which it was dosed into the process.

In brine coolers, the brine was in the working circuit. The brine was cooled to -20 ° C in an ammonia refrigeration system, and then it was fed to the appropriate processing elements. The brine was then homogenized in the brine tank and pumped back into the ammonia refrigeration system. The brine system included a compensation tank covered with oxygen-poor air (5 vol% oxygen in nitrogen).

The non-recyclable gas produced in the process passed through a separator and the non-condensed components were burned in a protected flare while the condensate was disposed of.

Obviously, the process can produce tert-butyl methacrylate of high purity (in this case, 99.68%), while at the same time energetically favorable removal of isobutene, which can be isolated with high separation efficiency from the esterification mixture, is carried out.

Example 1 - Option 1.

The process was carried out in accordance with example 1. Distillation purification was carried out using a natural circulation evaporator (4 bar abs. Steam, 114 pipes) without additional heating of the walls at the top.

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Temperature at the bottom of the distillation purification step: Temperature at the top of the distillation purification step:

approx. 102 ° C approx. 58 ° C approx. 53 ° C approx. 400 kg / h.

Wall temperature at the top:

Backflow ratio:

Steam heating for a natural circulation evaporator:

Column differential pressure: approx. 55 mbar

A study of the differential pressure in the purification column D2 was carried out as a function of time. Differential pressure and heating power are considered to be indicative of the degree of polymerization in the column. Over a period of about 8 days, an increase in the differential pressure in the purification column D2 was observed from about 55 mbar to about 110 mbar (FIG. 3). The reason for the increase in pressure was the condensation of gaseous products on the surfaces in the upper part of the column, these products formed unstabilized liquid phases subject to polymerization, which led to the appearance of deposits on the surfaces in the upper part of the column.

Also, a decrease in heating efficiency (heat transfer) was observed in the natural circulation evaporator. The required amount of steam used for heating was increased from about 400 to about 500 kg / h.

The upper two trays of the column were removed and required cleaning as they were covered with polymer pieces. Approximately 30% of the inner tubes of the evaporator were clogged with polymer and also required costly and effortless cleaning.

Example 1 - Option 2.

The process was carried out in accordance with example 1. Distillation purification was carried out using a natural circulation evaporator (4 bar abs. Steam, 114 tubes) with wall heating at the top with a half-closed tube shell (110 ° C, 1.5 bar).

Temperature at the bottom of the distillation purification step:

Temperature at the top of the distillation purification step:

Wall temperature at the top:

Backflow ratio:

Steam heating for a natural circulation evaporator:

Column differential pressure: approx. 55 mbar

A study was made of the heating power of the evaporator with natural circulation as a function of time. The heating power is considered to be an indication of the degree of polymerization in the column. In a natural circulation evaporator, a decrease in heating (heat transfer) power was observed. The required amount of steam used for heating was increased from about 400 to about 500 kg / h.

After 15 days, the process had to be interrupted. The top of the column was opened and no polymer deposits were found at the top of the column or on the top tray. However, approximately 30% of the inner tubes of the evaporator were clogged with polymer and required costly and effortless cleaning.

Example 1 - Option 3.

The process was carried out in accordance with example 1. Distillation purification was carried out using a forced circulation evaporator (4 bar abs. Steam, 308 tubes) with wall heating at the top with a half-closed tube shell (110 ° C, 1.5 bar).

Temperature at the bottom of the distillation purification step:

approximately 102 ° C approximately 58 ° C approximately 66 ° C 3.1 approximately 400 kg / h.

approx. 102 ° C approx. 58 ° C approx. 66 ° C HR approx. 400 kg / h.

Temperature at the top of the distillation purification step: Wall temperature at the top: Reflux ratio: Steam heating for a natural circulation evaporator:

Column differential pressure: approx. 55 mbar

A study was made of the differential pressure in the purification column D2, as well as the heating power of the forced circulation evaporator as a function of time. Differential pressure and heating power are considered to be indicative of the degree of polymerization in the column. No increase in differential pressure in the purification column was observed over a period of 30 days.

The process was interrupted. The top of the column was opened, no polymer deposits were found at the top of the column, nor on the top tray, nor in the forced circulation evaporator tubes.

Claims (14)

CLAIM 1. A method of continuous production of tert-butyl ester of ethylenically unsaturated carboxylic acid, comprising the following steps: a) the reaction of ethylenically unsaturated carboxylic acid with isobutene in the presence of an acid catalyst to obtain an esterification mixture; b) removing the acid catalyst; c) removal of low boiling point components; and d) feeding the liquid containing tert-butyl ether to a distiller, where it is purified by distillation, while: d1) in a still, separating the tert-butyl ether-containing liquid into a gaseous overhead product containing tert-butyl ether and a carboxylic acid-containing liquid underflow; d2) the gaseous overhead product containing tert-butyl ether is at least partially condensed and the condensate is at least partially recycled to the distiller as a reflux stream; d3) the carboxylic acid-containing liquid underflow is at least partially recycled to step a); d4) the carboxylic acid-containing liquid underflow is withdrawn and fed to a heating device; the superheated liquid recirculation stream is withdrawn from the heating device, which is released into the distiller; and d5) at least in the upper part of the distiller, the walls of the distiller in contact with steam in at least some areas are heated and / or thermally insulated. 2. A process according to claim 1, characterized in that the ethylenically unsaturated carboxylic acid is selected from acrylic and methacrylic acid. 3. A process according to any of the preceding claims, characterized in that the acetic acid content of the ethylenically unsaturated carboxylic acid is less than 300 ppm and the propionic acid content is less than 300 ppm. 4. A method according to any of the preceding claims, characterized in that: d1 ') the liquid underflow additionally passes through the device, in which it is purified from solid impurities. 5. The method of claim 4, wherein the device is a filter. 6. A method according to any of the preceding claims, characterized in that the heating device is a shell-and-tube heat exchanger operating in the mode of indirect heat exchange with the heating medium. 7. A method according to any of the preceding claims, characterized in that the superheated recirculation stream is released through a flow restrictor selected from a flow energy absorber, a valve, a converging nozzle, a perforated plate, a nozzle, a capillary tube, or combinations of these devices. 8. A process according to any one of the preceding claims, wherein the acid catalyst is an inorganic acid. 9. A method according to any one of claims 1 to 7, characterized in that the acid catalyst is an organic acid. 10. A method according to any of the preceding claims, characterized in that the content of the acid catalyst in the esterification mixture is 0.1-10% by weight. 11. A method according to any of the preceding claims, characterized in that the reaction in step a) is carried out in the presence of a stabilizer selected from phenothiazines. 12. A method according to any one of the preceding claims, characterized in that the stabilizer is selected from phenol compounds, the stabilizer being fed to the rectification section of the still. 13. A method according to any of the preceding claims, characterized in that the stabilizer is selected from N-oxyl compounds, the stabilizer being fed to the inlet of the still. 14. The method according to any of the preceding claims, characterized in that the volume ratio of oxygen to nitrogen in all gas mixtures used in steps a) to d) is in the range of 0.030.11.